Catalyzed alkylation, alkylation catalysts, and methods of making alkylation catalysts

Abstract

Improved alkylation catalysts, alkylation methods, and methods of making alkylation catalysts are described. The alkylation method comprises reaction over a solid acid, zeolite-based catalyst and can be conducted for relatively long periods at steady state conditions. The alkylation catalyst comprises a crystalline zeolite structure, a Si/Al molar ratio of 20 or less, less than 0.5 weight percent alkali metals, and further having a characteristic catalyst life property. Some catalysts may contain rare earth elements in the range of 10 to 35 wt %. One method of making a catalyst includes a calcination step following exchange of the rare earth element(s) conducted at a temperature of at least 575 C. to stabilize the resulting structure followed by an deammoniation treatment. An improved method of deammoniation uses low temperature oxidation.

Claims

1. A solid-acid alkylation catalyst, comprising: a zeolite structure comprising sodalite cages and supercages, a Si/Al molar ratio of 20 or less, less than 0.5 weight percent alkali metals, rare earth elements in the range of 10 to 35 wt %; and characterizable by a Catalyst Lifetime of 2 or greater where the Catalyst Lifetime parameter is defined as the catalyst age when olefin conversion falls below 90% using a test reaction where the solid-acid catalyst is loaded in a fixed-bed reactor such that the dT/dP>10 (diameter of tube to diameter of catalyst particles) and L/dP>50 (length of catalyst bed to diameter of catalyst particles) and exposed to a flow comprising a feed of 10:1 molar ratio of isobutane:n-butenes at 60 C. and 300 psig with a recycle ratio (R=volumetric flow rate of recycle stream/volumetric flow rate of feed stream) of 50, wherein the n-butenes comprise the olefin, where VS/VC is 7 (the ratio of system volume to catalyst volume), without regeneration of the solid-acid alkylation catalyst, and wherein a product of the test reaction has a RON of at least 92.

2. The solid-acid alkylation catalyst of claim 1 wherein the catalyst comprises 0.1 wt % to 5 wt % of an element selected from the group consisting of Pt, Pd, Ni, and combinations thereof.

3. The solid-acid alkylation catalyst of claim 1 having a catalyst lifetime of between 2.5 and 3.5.

4. The solid-acid alkylation catalyst of claim 1 where the Catalyst Lifetime parameter is defined as the catalyst age when olefin conversion falls below 95% using a test reaction where the solid-acid catalyst is loaded in a fixed-bed reactor such that the dT/dP>10 (diameter of tube to diameter of catalyst particles) and L/dP>50 (length of catalyst bed to diameter of catalyst particles) and exposed to a flow comprising a feed of 10:1 molar ratio of isobutane:n-butenes at 60 C. and 300 psig with a recycle ratio (R=volumetric flow rate of recycle stream/volumetric flow rate of feed stream) of 50, wherein the n-butenes comprise the olefin, where VS/VC is 7 (the ratio of system volume to catalyst volume), without regeneration of the solid-acid alkylation catalyst, and wherein a product of the RON of a product of the test reaction has a product is at least 92.

5. The solid-acid alkylation catalyst of claim 1 wherein the Catalyst Lifetime is 2.5 or greater.

6. A solid-acid alkylation catalyst, comprising: a zeolite structure comprising sodalite cages and supercages, a Si/Al molar ratio of 20 or less, less than 0.5 weight percent alkali metals, rare earth elements wherein the molar ratio of rare earth elements to (Si and Al) is in the range of 0.06 to 0.20; and characterizable by a Catalyst Lifetime of 2 or greater where the Catalyst Lifetime parameter is defined as the catalyst age when olefin conversion falls below 90% using a test reaction where the solid-acid catalyst is loaded in a fixed-bed reactor such that the dT/dP>10 (diameter of tube to diameter of catalyst particles) and L/dP>50 (length of catalyst bed to diameter of catalyst particles) and exposed to a flow comprising a feed stream comprising 100:1 molar ratio of isobutane:n-butenes, wherein the n-butenes comprise the olefin, at 60 C. and 300 psig without regeneration of the solid-acid alkylation catalyst, and wherein a product of the test reaction has a RON of at least 92.

7. The solid-acid alkylation catalyst of claim 6 wherein the catalyst comprises 0.1 wt % to 5 wt % of an element selected from the group consisting of Pt, Pd, Ni, and combinations thereof.

8. The solid-acid alkylation catalyst of claim 6 having a catalyst lifetime of between 2.5 and 3.5.

9. The solid-acid alkylation catalyst of claim 6 wherein the Catalyst Lifetime is 2.5 or greater.

Description

BRIEF DESCRIPTION OF THE DRAWINGS

(1) FIG. 1: Faujasite Structure of zeolite X and Y

(2) FIG. 2: Fixed-Bed Reactor configurations suitable for solid-acid alkylation

(3) FIG. 3: Fixed-bed micro-reactor for catalyst tests

(4) FIG. 4: Detailed reactor dimensions for catalyst tests. L is the length of the catalyst bed, dT is the tube diameter and dP is the catalyst particle diameter.

(5) FIG. 5: Representative solid-acid alkylation experimental results for Catalyst A.

(6) FIG. 6: Catalyst lifetime for lanthanum exchanged X zeolite calcined at different temperatures (catalysts A-D).

(7) FIG. 7: Catalyst lifetime for Y zeolite with (E, top line) and without (G, bottom line) a lanthanum exchange and subsequent calcination.

(8) FIG. 8: Catalyst lifetime for lanthanum exchanged X zeolite deammoniated with different moisture levels of <2 ppm (catalyst D) and 1.2% by vol. (catalyst H).

(9) FIG. 9: Catalyst lifetime for lanthanum exchanged X zeolite deammoniated at different temperatures (catalysts D and I-L).

(10) FIG. 10: Catalyst lifetime for lanthanum exchanged X zeolite at varying La+3 concentrations.

(11) FIG. 11: Catalyst lifetime for zeolite deammoniated at different temperatures (catalysts M-P).

(12) FIG. 12: Single stage feed recycle reactor schematic.

(13) FIG. 13: Representative recycle solid-acid alkylation experimental results for Catalyst D.

(14) FIG. 14: Representative Octane numbers for recycle solid-acid alkylation using Catalyst D.

(15) FIG. 15: Long term stability data results for Catalyst W.

(16) FIG. 16: Long term stability data results for Catalyst X.

(17) FIG. 17 shows catalyst lifetime as a function of deammoniation temperature and SAR for zeolite catalysts.

(18) FIG. 18. Butene conversion and product octane as a function of catalyst age for the Y zeolite deammoniated at 400 C. (Catalyst F).

DETAILED DESCRIPTION OF THE INVENTION

(19) The invention relies on a solid acid, crystalline zeolite structure that has both supercages and sodalite cages. These structures are well known and are shown in FIG. 1. It is believed that Bronsted acid sites in the supercages are important in catalyzing the alkylation reaction. The Faujasite structure with supercages and sodalite cages is known in Zeolite X and zeolite Y. Commercial zeolites are typically mixed with binders. The purity of zeolites can be measured by oxygen adsorption as described in Experimental Methods in Catalytic Research, vol. 2, Anderson and Dawson eds. (1976). Thus, although the invention can operate with pure crystalline zeolite, it typically operates with crystalline zeolite mixed with binder or other materials. Relatively high levels of aluminum are desirable; thus it is desirable to use zeolites with a Si/Al molar ratio of 20 or less, preferably 15 or less, in some embodiments in the range of 1 to 12, in some embodiments in the range of 2 to 10. We have found excellent results based on Zeolite X and acceptable performance with zeolite Y; nonetheless, based on the similarity of the pore structure, it is believed that zeolite Y should also function as a suitable starting material for the catalyst of the invention. The pore size of the sodalite cage structure does not exceed about 8 Angstroms, and the supercage structure has pores in which the pore size is at least about 10 Angstroms. In some preferred embodiments, the Si/Al molar ratio of zeolite X is in the range of 1 to 2; the Si/Al molar ratio of zeolite Y is in the range of 2 to 5; the Si/Al molar ratio of beta zeolite is in the range of 8 to 13.

(20) In a method of making a catalyst according to the present invention, a material containing a crystalline zeolite structure comprising sodalite cages and supercages and having a Si/Al molar ratio of 20 or less is treated with a solution containing a rare earth metal. The crystalline zeolite structure contains an alkali metal, typically sodium or potassium, most typically, sodium. The amount of alkali metal in the starting material is typically greater than 1 wt %, in some preferred embodiments greater than 3 wt %, in some embodiments between 5 and 20 wt %. The solution containing a rare earth metal is typically an aqueous solution. Preferred rare earth metals comprise lanthanum, cerium, neodymium, and praseodymium, and mixtures thereof; most preferably comprise lanthanum (La), and in some preferred embodiments the rare earth metal is at least 90% La or at least 95% La (by weight relative to total weight of all rare earth metals in solution). Preferably, the zeolite is treated with the rare earth solution at elevated temperature, preferably from 60 to 95 C., more preferably 70 to 90 C.; typically with a nitrate or sulfate salt solution. The solution containing a rare earth metal preferably has a concentration in the range of 0.1 M to 1.0 M, in some embodiments in the range of 0.4 to 0.8 M. Multiple treatments, for example, 3 treatments are preferred. Each treatment is preferably conducted for at least one hour at the elevated temperature, in some embodiments between 1 and 4 hours.

(21) If there is excess solution containing rare earth metal, it can be removed by decanting or filtering. Optionally, after decanting or filtering, the treated zeolite can be dried at temperatures up to 100 C. The resulting material is believed to have rare earth metal located in the supercages, but not yet exchanged with the alkali metal in the sodalite cages.

(22) To effectuate exchange of alkali ions in the sodalite cages with the rare earth ions located in the supercages, the catalyst is calcined at a temperature of at least 575 C. Although it was reported that the amount of La.sup.+3 in the sodalite cages becomes constant at temperatures above 300 C. (Monsalve, Thesis Active Acid Sites in Zeolite Catalyzed Iso-butane/cis-2-butene Alkylation Chap. 3, p 4), we surprisingly found significantly improved results from calcining at a much higher temperature. Preferably, the calcining step is carried out at a temperature of 575 to 650 C. In some preferred embodiments, the zeolite is held at a temperature between about 90 and 110 C. for at least 10 minutes, preferably at least 50 minutes. The zeolite can be heated at any suitable temperature ramping rate; for example between 1 C./min to 10 C./min. It may be preferred to hold the temperature at an intermediate value, such as between 200 and 300 C. for 30 min or more. Preferably, the zeolite is maintained at a temperature of at least 575 C., preferably between 575 and 650 C., in some embodiments between 600 and 625 C., or from 575 to 600 C., for at least 50 minutes, preferably for at least about 100 minutes; in some embodiments for between 50 and 500 minutes, in some embodiments between 50 and 240 minutes. Preferably, the entire calcination process, including temperature ramping times, is completed within 1 day, or completed within 2 days. The calcination step is preferably carried out at a relatively low humidity, for example, in dry flowing air containing less than 1 mass % water, in some embodiments less than about 50 ppm water. We believe that the calcination step causes some and, preferably essentially all, of the alkali metal ions (usually Na.sup.+) in the sodalite cages to be replaced with the rare earth ions (preferably La+3) from the sodalite cages.

(23) After calcination, the calcined zeolite is cooled and treated with an ammonium solution. The solution preferably has an ammonia concentration in the range of 0.1 M to 1.0 M, in some embodiments in the range of 0.2 to 0.5 M. This can be repeated several times; for example, from 2 to 5 times. One preferred set of conditions for the ammonium treatment is a temperature of from 50 to 100 C. for 10 minutes to 4 hours or more; more preferably from 30 minutes to two hours. In some embodiments of the invention, there is no rare earth exchange step and the zeolite (typically zeolite Y; containing Na cations) can be treated by the ammoniation process described herein.

(24) Any excess solution can be removed by decanting or filtration. The ammonium-exchanged zeolite can be heated to drive off excess water, for example to 100 C. or 200 C.

(25) Prior to use as a catalyst, the zeolite is converted from its ammonium form to the hydrogen form by heating, preferably in an atmosphere having very little water; for example, 1 mass % or less, or 0.2 mass %, or 2 ppm or less of water. This deammoniation temperature is preferably in the range of 300 to 400 C., more preferably 350 to 400 C.

(26) Although the scope of the present invention is not to be limited to any theoretical reasoning, it is believed that the deammoniation step converts the ammonium cation sites to Bronsted acid sites, especially in the supercages, while the rare earth elements remain in the sodalite cages. Because the acid, or H+, sites are located in the larger diameter supercage structure of the catalyst, pore mouth plugging is significantly reduced, allowing the catalyst to remain active for increased periods of time, while the rare earth metal cation sites, such as, for example, La.sup.+3 sites, provide enhanced stability to the sodalite structure. We believe that at least 80% of the cationic sites in the sodalite portion are rare earth metal cation sites, and at least 80% of the cationic sites in the supercage portion are H+ sites.

(27) We have found that careful control of the deammoniation conditions for the zeolite catalyst lead to improvements in catalyst performance, when converting the ammonium form of the zeolite to the active or acid form. When the ammonium form of a zeolite is heated, the initial step is the evolution of physically adsorbed water, which causes a first endotherm at about 150 C.; this step is completed at 200 C. Ammonia then is evolved which gives rise to a second endotherm at 300 C.; this step is completed at about 400 C. Raising the temperature above 400 C. results in evolution of water from the condensation of the hydroxyl groups. This dehydroxylation step results in a) a significant decrease in the number of active catalytic acid sites and b) conversion of the preferred Bronsted acid sites to the Lewis acid sites which increases the rate of catalyst deactivation.

(28) The invention also relates to a reactor suitable for paraffin alkylation using solid acid catalysts. Paraffin alkylation is a fast reaction, which benefits from low olefin concentrations (typically the reactor I/O ratio>300) in the reactor to suppress the polymerization reaction. In conventional liquid-acid based reactors, high speed mechanical agitators are used to disperse the hydrocarbon feed into the acidic medium. Specially designed jets are used to introduce the olefin feed as small droplets to avoid high localized olefin concentration. A departure from perfect mixing conditions results in significant deterioration of product octane quality and formation of Acid Soluble Oils via olefin polymerization reaction which leads to higher acid consumption. The only way to achieve the same level of mixing with solid-catalysts, is to use a slurry system. However, slurry systems are difficult to handle and equipment needed to pump slurries around are very expensive.

(29) Fixed-bed reactors are easier to design, scale-up and maintain and, therefore, preferred embodiments utilize a fixed bed reactor. One way of achieving a low olefin concentration in the bulk liquid is obtained by staging the olefin feed over the catalyst bed. This approach is often used in designing reactors for aromatic alkylation reactions for the production of ethylbenzene or cumene. Typically 4-6 stages (FIG. 2A) are used in designing such reactors. In paraffin alkylation reactions, the feed mixture of iso-butane to olefins (I/O) range from 10-15 molar and a typical catalyst bed I/O target is 300-500. This implies that the olefin must be distributed in 30-50-stages to meet this recommended dilution in olefin concentration which is difficult to achieve in commercial scale reactors. Another way of achieving this level of olefin dilution without resorting to staging the feed is by introducing a recycle stream of the reactor effluent (FIG. 2B). By using a recycle ratio of 30-50, a feed with an I/O ratio of 10-15 would produce a catalyst bed I/O of 300-500. However, the large volumetric flow through the catalyst bed would generate a proportionally large pressure-drop greatly increasing the duty (hence power consumption) on the recycle pump. A novel way of designing a fixed-bed reactor for solid acid catalyzed reactions is to combine the two concepts into a hybrid reactor (FIG. 2C) where the olefin feed is staged over 4-6 zones and the recycle pump recycles the product stream at a modest recycle ratio of 6-10. This allows a catalyst bed I/O of 300-500 without resorting to a large number of feed stages or a very high recycle ratio. An added benefit of this design is the ability to remove the heat of reaction externally. The invention includes methods using the reactors described herein and includes systems (apparatus plus fluids and, optionally, conditions within the apparatus) that includes the reactors described herein.

(30) The invention is further elucidated in the examples below. In some preferred embodiments, the invention may be further characterized by any selected descriptions from the examples, for example, within 20% (or within 10%) of any of the values in any of the examples, tables or figures; however, the scope of the present invention, in its broader aspects, is not intended to be limited by these examples.

EXAMPLES

Example 1Catalyst A

(31) The starting material was a commercial zeolite X having a SiO.sub.2/Al.sub.2O.sub.3 molar ratio of 2.8 (Si/Al of 1.4) and a sodium content of 15% by weight. 5 grams of the zeolite was crushed and sieved to 0.5-1.4 mm particles. They were suspended in 50 mL of deionized water and stirred for 15 minutes after which the water was decanted. This washing procedure was repeated a second time.

(32) A lanthanum ion exchange was performed immediately following the initial water wash. The zeolite was suspended in 50 mL of a 0.8 M lanthanum nitrate solution and heated to 80 C. while stirring for 2 hours. The lanthanum solution was decanted and replaced with a fresh solution. This lanthanum exchange was performed three times followed by 2 water washes of 75 mL each. The zeolite was then left to dry at room temperature.

(33) Following the lanthanum exchange, the zeolite was calcined in a muffle furnace. The temperature program for calcination was 1.5 C./min ramp to 100 C. where it was held for 1 hour, 2.0 C./min ramp to 230 C. and hold for 2 hours, 10 C./min ramp to the final calcination temperature of 400 C. for 4 hours.

(34) The lanthanum exchanged zeolite was suspended in a 0.5 M ammonium nitrate solution and heated to 80 C. with stirring for 2 hours. The ammonium solution was decanted and replaced with fresh solution. This ion exchange was performed 3 times followed by 2 water washes of 75 mL each. The zeolite was then left to dry at room temperature. The zeolite was deammoniated in dry air (<2 ppm) using the following temperature program: 100 C. (0.5 hours), 120 C. (1 hour), 230 C. (2 hours), 400 C. (4 hours). 400 C. is the deammoniation temperature required to convert the catalyst from the ammonium form to the active proton form. The lower temperatures are necessary to completely dry the catalyst.

Example 2Catalyst B

(35) The catalyst was prepared as in Example 1 with the only difference being the final calcination temperature. In this example the final calcination temperature following lanthanum exchange was 450 C.

Example 3Catalyst C

(36) The catalyst was prepared as in Example 1 with the only difference being the final calcination temperature. In this example the final calcination temperature following lanthanum exchange was 550 C.

Example 4Catalyst D

(37) The catalyst was prepared as in Example 1 with the only difference being the final calcination temperature. In this example the final calcination temperature following lanthanum exchange was 600 C.

Example 5Catalyst E

(38) The catalyst was prepared as in Example 1. However, the starting material used was a Y zeolite in this example. The commercial Y zeolite had a SiO.sub.2/Al.sub.2O.sub.3 molar ratio of 5.0 and a sodium content of 14% by weight. Since the Y zeolite is in powder form it must be filtered rather than decanted in each solution exchange. Additionally, it is pelletized following ammonium exchange and drying then crushed and sieved to 0.5-1.4 mm catalyst particles.

Example 6Catalyst F

(39) The catalyst was prepared as in Example 5 with the only difference being that no Lanthanum exchange and subsequent calcination was performed. Following the initial water wash, the Y zeolite undergoes an ammonium exchange and deammoniation. In this example the deammoniation temperature was 400 C.

Example 7Catalyst G

(40) The catalyst was prepared as in Example 5 with the only difference being that no Lanthanum exchange and subsequent calcination was performed. Following the initial water wash, the Y zeolite undergoes an ammonium exchange and deammoniation. In this example the deammoniation temperature was 550 C.

Example 8Catalyst H

(41) The catalyst was prepared as in Example 3 with the only difference being water content of the air used for activation following ammonium exchange. In this example, the water content was 1.2% by volume.

Example 9Catalyst I

(42) The catalyst was prepared as in Example 3 with the only difference being the deammoniation temperature used following ammonium exchange. In this example the deammoniation temperature was 300 C.

Example 10Catalyst J

(43) The catalyst was prepared as in Example 3 with the only difference being the deammoniation temperature used following ammonium exchange. In this example the activation temperature was 350 C.

Example 11Catalyst K

(44) The catalyst was prepared as in Example 3 with the only difference being the deammoniation temperature used following ammonium exchange. In this example the deammoniation temperature was 450 C.

Example 12Catalyst L

(45) The catalyst was prepared as in Example 3 with the only difference being the deammoniation temperature used following ammonium exchange. In this example the deammoniation temperature was 500 C.

Example 13Catalyst M

(46) The catalyst was prepared as in Example 1. However, the starting material used was a zeolite in this example. The commercial zeolite had a SiO.sub.2/Al.sub.2O.sub.3 molar ratio of 16. The zeolite does not undergo a lanthanum exchange and the subsequent calcination. Following an initial water wash, it is immediately exchanged with ammonium 3 times. It is then deammoniated in dry air with a final temperature of 400 C.

Example 14Catalyst N

(47) The catalyst was prepared as in Example 13 with the only difference being the deammoniation temperature used following ammonium exchange. In this example the deammoniation temperature was 450 C.

Example 15Catalyst O

(48) The catalyst was prepared as in Example 13 with the only difference being the deammoniation temperature used following ammonium exchange. In this example, the deammoniation temperature was 500 C.

Example 16Catalyst P

(49) The catalyst was prepared as in Example 13 with the only difference being the deammoniation temperature used following ammonium exchange. In this example the deammoniation temperature was 550 C.

Example 17Catalyst Q

(50) The catalyst was prepared as in Example 13 with the only difference being the starting -zeolite had a SiO.sub.2/Al.sub.2O.sub.3 molar ratio of 25.

Example 18Catalyst R

(51) The catalyst was prepared as in Example 13 with the only difference being the starting -zeolite had a SiO.sub.2/Al.sub.2O.sub.3 molar ratio of 75.

Example 19Catalyst S

(52) The catalyst was prepared as in Example 17 with the only difference being the deammoniation temperature used following ammonium exchange. In this example the deammoniation temperature was 550 C.

Example 20Catalyst T

(53) The catalyst was prepared as in Example 18 with the only difference being the deammoniation temperature used following ammonium exchange. In this example the deammoniation temperature was 550 C.

Example 21Catalyst U

(54) The catalyst was prepared as in Example 4 with the only difference being the lanthanum ion exchange was performed using a 0.3 M lanthanum nitrate solution.

Example 22Catalyst V

(55) The catalyst was prepared as in Example 4 with the only difference being the lanthanum ion exchange was performed using a 0.5 M lanthanum nitrate solution.

Example 23Catalyst W

(56) The catalyst was prepared as in Example 4 with the only difference being the lanthanum ion exchange was performed using a 0.6 M lanthanum nitrate solution.

Example 24Catalyst X

(57) The catalyst was prepared as in Example 4 with the only difference being the lanthanum ion exchange was performed using a 0.8 M lanthanum nitrate solution.

Example 25Catalyst Y

(58) The catalyst was prepared as in Example 4 with the only difference being the lanthanum ion exchange was performed using a 1.0 M lanthanum nitrate solution.

Example 26Catalyst Z

(59) The catalyst was prepared as in Example 21. The catalyst was impregnated with Tetraamine Platinum Chloride to give 0.1 wt % Pt loading on the catalyst.

Example 27Catalyst AA

(60) The catalyst was prepared as in Example 24. The catalyst was impregnated with Nickel Nitrate to give 0.25 wt % Nickel loading on the catalyst.

(61) Alkylation activity experiments were performed using an isothermal packed bed reactor setup. Heating is controlled using an Omega temperature control unit and a ceramic heating element. Feeds are sent through a preheater of 75 cm length prior to entering the reactor.

(62) The catalyst of interest (1 g) is first loaded into a reactor shown in FIG. 3 (7.9 mm diameter), a center thermocouple (K-type) is inserted and positioned such that the tip of the thermocouple (3.1 mm diameter) is at the bottom of the catalyst bed. 1 mm glass beads are used to fill any void space in the reactor. The catalyst is deammoniated in dry air (GHSV=1000 hr.sup.1) at atmospheric pressure using the following temperature program: 100 C. (0.5 hour), 120 C. (1 hour), 230 C. (2 hours), 400 C. (4 hours) (these values are for Example 1). Following deammoniation the reactor is allowed to cool to reaction temperature (75 C.), then purged with dry nitrogen (GHSV=1000 hr.sup.1) for 0.5 hours. The reactor is pressurized (300 psig) with pure isobutane to begin the experiment.

(63) The reaction feed is contained in helium-purged Hoke cylinders. Isobutane and 1-butene (source for both is AGL Welding Supply Co, Ltd) are analyzed for any impurities using a HP5890 GC equipped with a Petrocol DH column. All feed and product analysis uses this GC system with the following program: 60 C. (16 min), ramp at 15 C./min to 245 C. and soak (20 min).

(64) The experiment is run using an olefin hourly space velocity equal to 0.5 hr.sup.1 and a feed I/O ratio of 100. This equates to 40 g/hr feed rate for isobutane and 0.4 g/hr for 1-butene. The flow rates are controlled by Eldex ReciPro Model A pumps. Product samples are extracted using a high pressure sampling port and syringe (Vici Precision Sampling) and immediately injected into the HP5890 GC for analysis.

(65) Regeneration may be performed using hydrogen gas (1000 hr.sup.1 GHSV) at a regeneration temperature of 250 C. for 2 hours. Process and detailed reactor schematics are shown in FIGS. 3 and 4.

Application Example 1

(66) The lanthanum exchanged X zeolites were prepared with different calcination temperatures as in Examples 1-4 (Catalyst A-D). 1 gram of each catalyst was loaded into a reactor shown in FIG. 3. The reactor was purged with nitrogen at a temperature of 75 C. It is then pressurized with isobutane to 300 psig. The reaction feed mixture, I/O Ratio of 100, was fed to the reactor with an OHSV of 0.5 hr.sup.1. Product samples were withdrawn periodically from a high pressure sample port and analyzed using a gas chromatograph equipped with a Petrocol DH 100 m column.

(67) FIG. 5 shows representative data using catalyst A from Example 1. Research Octane Number (RON) and Motor Octane Number (MON) are calculated for the alkylate from product analysis.

(68) FIG. 6 shows catalyst lifetime as a function of calcination temperature. The zeolite calcined at 600 C. (Catalyst D) possessed a lifetime of 3.75.

(69) As can be seen from FIG. 6, catalyst lifetime plateaus between 450 and 550 C.an observation which matches the conventional understanding that the zeolite catalyst be calcined at a temperature of 450 C. or less. Surprisingly, however, we discovered that heating to 575 C. or higher, more preferably 600 C. resulted in a substantial improvement in catalyst lifetime, and preferably less than 650 C. to avoid degradation of the zeolite structure.

Application Example 2

(70) The Y zeolites were prepared with and without lanthanum exchange steps followed by calcination as in Examples 5 and 7 (Catalysts E-G). The experimental conditions are identical to those of Application Example 1.

(71) FIG. 7 shows butene conversion as a function of catalyst age for the Y zeolites. As shown, the catalyst age of the H form of zeolite Y is substantially increased by lanthanum exchange followed by calcination.

Application Example 3

(72) The catalysts used were catalyst D (<2 ppm) and catalyst H (Example 8, 1.2% by volume) at different water contents in the air during deammoniation. The experiment is identical to Application Example 1.

(73) FIG. 8 shows catalyst lifetime for both the water concentrations used during deammoniation. Moisture in the oxidizing gas (typically air) during deammoniation should be avoided. Therefore, the atmosphere (or, more typically, the gas that flows over the zeolite as measured prior to contacting the zeolite) during deammoniation preferably comprises 0.2 vol % or less, more preferably 10 ppm or less, more preferably 5 ppm or less, and still more preferably 2 ppm or less water.

Application Example 4

(74) The catalysts used were from examples 4 (catalyst D) and 9-12 (catalysts I-L). They were deammoniated at different temperatures under dry conditions (<2 ppm). FIG. 9 compares lifetime for the catalysts deammoniated at different temperatures. There is a maximum catalyst lifetime of 2.9 at 400 C. deammoniation temperature. The deammoniation temperature was maintained at 400 C. for 4 hours.

(75) The superior catalyst lifetime results for deammoniation in the range of about 400 to 450 C. was especially surprising since the guidelines from Linde Molecular SievesCatalyst Bulletin, Ion-Exchange and Metal Loading Procedures state that to decationize NH.sub.4.sup.+ exchanged molecular sieve should be conducted in dry air at 550 C. for 3-4 hours.

Application Example 5

(76) The catalysts used were from examples 18-22 (catalysts RV). They were deammoniated at 400 C. under dry conditions (<2 ppm). FIG. 10 compares lifetime for the catalysts ion-exchanged with varying concentrations of Lanthanum Nitrate solutions. There is a maximum catalyst lifetime of 3.7 at 0.8 M Lanthanum Nitrate concentration. FIG. 10 shows that catalyst lifetime reaches a maximum, in a single exchange step, where La concentration is in the range of about 0.5 to 0.9 M, preferably 0.6 to 0.8 M; La concentrations above this range lowers pH and thus causes structural collapse.

Application Example 6

(77) The zeolites were prepared with different deammoniation temperatures as in Examples 13-16 (Catalyst M-P) and loaded into a fixed-bed reactor. In this experiment the reaction was run in recycle mode. The reaction feed mixture, I/O Ratio of 15, was fed to the reactor at a rate of 10 g/hr. The recycle stream flow rate was 40 g/hr. The combined feed rate to the reactor was 50 g/hr with an OHSV of 0.2 hr.sup.1. Product samples were withdrawn periodically from a high pressure sample port and analyzed using a gas chromatograph equipped with a Petrocol DH 100 m column as in Application Example 1

(78) FIG. 11 shows catalyst lifetime as a function of deammoniation temperature for these catalysts. The highest performance is for the catalyst deammoniated at 400 C. with dry air. Thus, we observed the surprising result that the deammoniation temperature (450 C. to 40 C., preferably 425 C. to 40 C.) resulted in superior catalyst lifetimes. This was a very surprising result since it had been reported that activation at 45 C. under dry conditions resulted in barely active catalysts and that a temperature of 55 C. was required for significantly improved activity. See Kunkeler et al., Zeolite Beta: The Relationship between Calcination Procedure, Aluminum Configuration, and Lewis Acidity

Application Example 7

(79) The lanthanum exchanged X zeolite from Example 4 (Catalyst D) was loaded into a fixed-bed reactor with product recycle shown in FIG. 12. A reaction feed mixture, with a I/O molar ratio of 10 (with n-butene as olefin), was fed to a bench-scale reactor with an OHSV of 0.1 hr.sup.1 at a temperature of 60 C. and a pressure of 300 psig. This recycle flow rate was established such that the recycle ratio was 50. Product samples were withdrawn periodically from a high pressure sample port and analyzed using a gas chromatograph equipped with a Petrocol DH 100 m column.

(80) FIG. 13 shows representative product distribution data using catalyst D from Example 4.

(81) FIG. 14 shows reaction octane numbers as a function of run time. Research Octane Number (RON) and Motor Octane Number (MON) are calculated for the alkylate from product analysis.

(82) The lifetime of this catalyst was >3.25 under commercial reaction conditions before regeneration. The steady state product C.sub.8 selectivity was 79 wt %, RON was 97 and the product MON was 93.

Application Example 8

(83) The lanthanum exchanged X zeolite from Example 23 (Catalyst Z) was loaded into a fixed-bed reactor with product recycle shown in FIG. 15. A reaction feed mixture, with a I/O molar ratio of 10 (with MTBE raffinate as olefin), was fed to a bench-scale reactor with an OHSV of 0.1 hr.sup.1 at a temperature of 45 C. and a pressure of 300 psig. The catalyst test was run for 24 hours and then regenerated with hydrogen gas at 250 C. for 2 hours. This cycle was repeated for 18 months. Results of this test are shown in FIG. 15.

(84) Data shown in FIG. 15 demonstrates that 0.1 wt % Platinum loading on the catalyst is adequate for regeneration the catalyst with hydrogen gas.

Application Example 9

(85) The lanthanum exchanged X zeolite from Example 27 (Catalyst AA) was loaded into a fixed-bed reactor with product recycle shown in FIG. 16. A reaction feed mixture, with a I/O molar ratio of 10 (with MTBE raffinate as olefin), was fed to a bench-scale reactor with an OHSV of 0.1 hr-1 at a temperature of 45 C. and a pressure of 300 psig. The catalyst test was run for 24 hours and then regenerated with hydrogen gas at 250 C. for 2 hours. This cycle was repeated for 50 days.

(86) The data shown in FIG. 16 demonstrates that about 0.25 wt % Nickel loading on the catalyst is adequate for regeneration the catalyst with hydrogen gas; a preferred range is 0.1 to 0.5 wt % Ni.

Application Example 10

(87) The zeolites were prepared with different Silica-to-Alumina Ratios (SAR) and deammoniation temperatures as in Examples 17-20 (Catalysts Q-T) and loaded into a fixed-bed reactor with product recycle shown in FIG. 14. The reaction feed mixture, I/O Ratio of 15, was fed to the reactor at a rate of 10 g/hr. The recycle stream flow rate was 40 g/hr. The combined feed rate to the reactor was 50 g/hr with an OHSV of 0.2 hr-1. Product samples were withdrawn periodically from a high pressure sample port and analyzed using a gas chromatograph equipped with a Petrocol DH 100 m column as in Application Example 1.

(88) FIG. 17 shows catalyst lifetime as a function of deammoniation temperature and SAR for these catalysts. The highest performance is for the catalyst at SAR 16 and deammoniated at 400 C. Results shown in FIG. 17 are unique in the way the NH4+ form is converted to the H+ form. The typical deammoniation temperature used for zeolites is typically 550 C. or higher. Our results clearly show a much superior performance at lower deammoniation temperatures.

Application Example 11

(89) The Y zeolites were prepared without Lanthanum exchange steps followed by deammoniation as in Examples 6 (Catalysts F). The experimental conditions are identical to those of Application Example 1

(90) FIG. 18 shows butene conversion and product octane as a function of catalyst age for the Y zeolite deammoniated at 400 C. (Catalyst F)showing superior catalyst lifetime for deammoniation at this temperature.

(91) Comparing performance of Y-zeolite deammoniated at 400 C. (Catalyst F) with Y-zeolite deammoniated at 550 C. (catalyst G) clearly demonstrates the superiority of the low temperature deammoniation method.

(92) It is to be understood, however, that the scope of the present invention is not to be limited to the specific embodiments described above. The invention may be practiced other than as particularly described and still be within the scope of the accompanying claims.